Hydrocarbon conversion process



July 23, 1968 J. ENG ET AL HYDROCARBON CONVERSION PROCESS 2 Sheets-Sheet l Filed Feb. 20, 1961 NNE wNz July 23, 1968 J. ENG x-:T AL

HYDROCARBON CONVERSION PROCESS 2 Sheets-Sheet 2 Filed F'eb. 20, 1961 United States Patent O 3,394,199 HYDROCARBON CONVERSON PROCESS Jackson Eng and Roger M. Butler, Sarnia, Ontario,

Canada, assignors to Esso Reseach and Engineering Company, a corporation of Delaware Filed Feb. 20, 1961, Ser. No. 90,408 14 Claims. (Cl. 2611-674) The present invention pertains to an improved hydrocarbon conversion process and, more particularly, to an improved method for the production of benzene.

Steam cracking of hydrocarbon feedstocks is a well known and widely used process for the production of petrochemicals, particularly ethylene, propylene, and butadiene. A variety of feedstocks boiling up to and including the gas oil boiling range can be used, although usually a light gasoline or naphtha charge stock is used. The steam cracked product comprises, in addition to ethylene, propylene, and butadiene, a gasoline fraction and a residue or fuel oil fraction. The gasoline fraction has a Research Octane Number of about 95 to 100` and is, therefore, useful as a motor gasoline blending component. The steam cracked naphthas consist chiefly of aromatics and olens, and the 140-200" F. fraction is especialy rich in benzene (approximately 40-70 vol. percent), and, therefore, a potential feedstock to an aromatics extraction plant. Glycol-type solvents are conventionally used for the extraction of aromatics from hydrocarbon fractions. Unfortunately, glycol-type solvents do not give a sharp separation of aromatics from oletins and particularly dioletins. A further difficulty is that appreciable quantities of thiophene occur in some steam cracked naphthas. This is extracted with the benzene resulting in an impure product. It is the object of this invention to provide an improved method for preparing aromatic hydrocarbons and, in particular, benzene.

It is also an object of this invention to provide an improved method for extracting aromatics from steam cracked naphthas.

It is a further object of this invention to provide a method for treating steam cracked naphthas to render them more amenable to extraction with glycol-type solvents. I

These and other objects will appear more clearly from the detailed specification and claims which follow. p,

It has now been found that aromatic hydrocarbons can be eiciently extracted from steam cracked naphthas by means of glycol-type solvents, provided that the naphtha fraction is freed of `diolelin or other polyunsaturates, and substantially freed of monoolens prior to the extraction treatment.

In accordance with the present invention, the dioleiins and monoolens in steam cracked naphthas are saturated in an efficient twostage hydrogen treatment. In the rst stage, the naphtha is hydroned with a reduced nickel catalyst in order to largely eliminate the dioleiins and thereby stabilize the naphtha stream. Because of the high activity of the nickel metal catalyst, operations can be carried out at about 300 F. At this low temperature, polymer and coke formation are avoided and, consequently long catalyst life is possible. After the first stage treatment, the stabilized naphtha is hydroned with the usual cobalt oxide-molybdenum oxide type catalyst. While this second stage treatment requires temperatures of about 500 F. or higher, there is no polymer and coke formation problem in substantially completely saturating the rest of the monoolelins in this stage because the unstable dioleins have been removed. The treated product from the second stage is completely satisfactory for charging to the solvent extraction unit.

Reference is made to the accompanying drawings, (FGURES 1 and 2) which illustrate diagrammatic flowplans of two embodiments of the process of the present invention.

Referring to FIGURE l, a naphtha or gas oil feed is introduced through line 1 and mixed with the steam supplied through line 2 and enters cracking furnace 4 via line 3. The quantity of steam added is controlled so as to make the concentration of steam within the furnace from 50 to 9() mole percent, preferably 80 mole percent. ln the cracking furnace, the temperature is raised to 1300 to 1450 F., preferably 1400 to 1440. The pressure within the cracking furnace is normally in the range of 10 to 20 p.s.i.g. The residence time in the furnace at cracking temperature is 0.2 to 2 seconds. The cracked products are discharged from the furnace via line 5 to separation equipment 6. This equipment is of a conventional nature and involves fractionation and, in some cases, extraction. The steam cracked naphtha is withdrawn from the separator, via line 8. Other steam cracked products leave via lines 7. The steam cracked naphtha is in the boiling range of C5 to 400, preferably C5 to 300. The naphtha in line 8 goes to the caustic treater 9 in which it is given a mild treat to remove mercaptans. Typical conditions would consist of a 10% treat with 20 B. caustic. The mercaptan-free naphtha leaves the treater via line 10 and is mixed with a stream of hydrogen sulfide-free, hydrogen-rich, treat gas., which enters via line 11. This gas will normally consist of caustic washed reformer tail gas and will have a concentration of about mole percent hydrogen. Normally the quantity of treat gas will be in the range of 750 to 2000 standard cu. ft./ bbl. of fresh feed. The mixed naphtha and hydrogen proceeds through heat exchanger 12, then via lines 13 and 1S into the first stage reactor 16. Recycled product is mixed with the naphtha via line 14. The heat added in exchanger 12 is controlled so that the feed entering the reactor is in the range of Z50-300 F. Normally, low temperatures will be required with fresh catalyst and the temperature may be raised to 'the upper limit as the catalyst deactivates. The catalyst in the first stage reactor 16 is a reduced nickel type catalyst.

This may consist of the familiar type of catalyst known as nickel supported on kieselguhr. Such catalysts normally contain about 10 to 50 weight percent nickel, prefera bly 40 to 50. The preparation of a suitable nickel-onkieselguhr catalyst has -been described by L. W. Covert, R. Connor and H. Adkins (I.A.C.S., 54, 1951 (1932)). Catalysts of this type may be used in granular, pelleted or extruded form. Cylindrical tablets one-eighth inch in diameter by one-eighth inch tall are satisfactory although many other sizes can also be used.

Other types of supported nickel catalyst which can be used consist of nickel on precipiated calcium carbonate and nickel on alumina. The pressure in 16 is in the range of 150-3-00 p.s.i.g. and the liquid hourly space velocity is in the range of .5 to 4 v./v.hr. based on fresh feed. The product from the reactor proceeds via line 17 into separator 18. Here the gas phase leaves via line 20 and the liquid phase via line 19. Normally, the separator will be at a higher temperature than the feed inlet because of the considerable exothermic heat of reaction This is controlled by recycling product via line 21 and pump 22 and line 14. The quantity of this recycle will normally be in the volume of 3 to 10 times that of the fresh feed. It should be chosen such that the temperature rise in the rst stage reactor is less than 100 F., preferably less than 50 F. The liquid product is pumped via line 23 and pump 24, through heater 25 line 26 and line 30 into the second stage reactor 31. Hydrogen-rich gas from the rst stage separator 18 passes through line 20, is compressed by compressor 27, and passes through line 30. Recycle hydrogen and fresh hydrogen-rich treat gas also mix with the feed via line 29. The second stage reactor contains a sulded cobalt molybdate-on-alumina catalyst. The second stage hydrotreat vessel is charged with a conventional metal oxide-type hydrogenation catalyst, prefera bly a cobalt molybdate catalyst or one containing cobalt oxide and molybdenum oxide dispersed upon an aluminacontaining support or base, preferably activated or adsorptive alumina. In general, such catalysts are prepared by first forming adsorptive alumina particles or hydrous alumina which on heating becomes adsorptive, and then compositing molybdenum oxide and cobalt oxide therewith. The molybdenum oxide can, for example, be added as a slurry or it may be applied as a solution of ammonium molybdate. The cobalt oxide is conveniently added as a salt such as cobalt nitrate or acetate, salts which are readily decomposed to cobalt oxide and volatile materials. The cobalt oxide and molybdenum oxide may be provided in equimolar amounts or a molecular excess of one over the other may lbe used. Suitable catalysts contain from about 5 to about 25 wt. percent cobalt oxide and molybdenum oxide, with the ratio of the former to the latter in the range of from about 1/5 to about 5/1. The pressure in this reactor is in the range of 400 to 1000 p.s.i.g., `preferably 500 to 800 p.s.i.g. The inlet temperature is controlled by means of heater 25 to be within the range of 550 to 600, preferably about 575. Again there is a large exothermic heat of reaction and in the process shown in FIGURE 1, this is controlled by the addition of recycled product at intermediate points within the reactor. This cold quench will be at a temperature of 50 to 150 F. and will be in the liquid phase. As it enters the reactor, considerable vaporization occurs thus absorbing the heat of the reaction. The outlet temperatures from the reactor are maintained within the range of 550 to 700 F. and preferably within the range of S75-675 F. The liquid hourly space velocity in the second stage is in the range of .5 to 4 V./v.-hr. based on fresh feed. The product leaves the reactor via line 32 and is cooled and partially condensed in cooler 33. If the proceeds to separator 35 via line 34. The liquid product leaves the separator via line 42 and is divided into recycle which proceeds through line 44 to inlets 45 at reactor 31 and product which goes through line 43. The gas from the separator proceeds through line 36 and is partially recycled via line 37, compressor 38 and line 39. The remainder of the gas leaves the unit via line 41. The total hydrogen entering the second stage reactor is controlled so that it is in the range of 1000 to 3000 standard cu. ft. of hydrogen/bbl. of total liquid feed. The concentration of hydrogen in the product line 32 should be a minimum of 40 mole percent. The doubly hydrofined product enters the depentanizer 46 via line 43. Here the pentanes and lighter components are taken overhead via line 47 and the higher lboiling fraction leaves via line 48 to the extraction unit 49.

This extraction unit may employ any conventional solvent such as diethylene glycol-water mixtures, triethylene glycol-water mixtures, dipropylene glycol-water mixtures or combinations of these solvents, sulfur dioxide, the sulfolanes, etc. The design and operation of such units is well known in the art. The raffinate consisting essentially of saturated hydrocarbons leaves the extraction unit via line 50. The aromatic extract is pumped via line 51 through heater 52. Here it is raised to a temperature of the order of 1GO-600 F., preferably 500 F. at a pressure of 200 to 500 p.s.i.g. The heated aromatic extract passes via line 53 into clay treating drum 54. This contains active clay particles such as, for example, that known as attapulgife. The quantity of clay in drum 54 is chosen so as to make the liquid hourly space velocity about .5 to 1 v./v.hr. The purpose of the clay treating step is to remove any traces of olens and diolefins which may be present. The clay treated aromatic extract passes via line 55 to the benzene tower 56 in which an essentially pure benzene product is removed via line 57. The bottoms from tower 66 passes through line 58 to tower 59 where essentially pure toluene is removed overhead in line 60 and heavier aromatics in line 61.

FIGURE 2 shows a variation of the process presented in FIGURE 1. In this scheme, a gasoline fractionator is incorporated between the rst and second stage hydroners. The purpose of this fractionator is to remove 140- and 200/400 streams that are suitable for incorporation into motor gasolines without further treatment. The process is exactly the same up to and including separator 1S. The rst stage hydroner product from the separator passes through line 70 into the fractionator 71. Here a 140-200 F. boiling range benzene heart cut is prepared. This passes via line 74 into second stage feed pump 75. The 140- overhead stream and the 200-400 stream from the fractionator may be used as motor gasoline blending components. The 400+ fraction from the fractionator is highly `aromatic and is unsuitable for gasoline or heating oil. It may be used either as a naphthalene concentrate or incorporated in cat cracker feed or bunker fuel. The second stage feed passes from the feed pump 75 through line 76 through heater 77 and line 78 and is mixed with hydrogen-rich treat gas which enters via line 91. The combined feed and gas enter the second stage reactor 80 via line 79 at a pressure of 400- 1000 p.s.i.g., preferably 50G-800 p.s.i.g. The inlet temperature is controlled by means of heater 77 to be within the range yof 550 to 600 F., preferably about 575 F. The exothermic heat of reaction is controlled in the process shown in FIGURE 2 by means of indirect internal reactor cooling, instead of the cooled quench injection shown in FIGURE 1. This is merely an alternative method. The design of these coolers should be such that the reactor temperature does not rise above 700 F. and preferably not above 675. The reactor products leave via line 81 and are cooled and partially condensed in heat exchanger 104. The liquid is separated from the gas in separator S2. The gas leaves via line 83 and is partially recycled via line 87 and line 91. Tail gas may, if it is desired, pass to a hydrogen purification plant. The liquid fr-om the separator 82 passes via line 92 into the extraction unit 93. This extraction unit may employ any of the conventional solvents described in the discussion of FIGURE 1; its purpose is to separate the feed into an essentially saturated raffinate which leaves via line 94 and a saturate-free benzene concentrate which leaves via line 95. This is clay treated by means of heater 96 line 97 and clay drum 98 at conditions similar to those used in FIGURE 1. The clay treated product passes through line 99 into the distillation tower 100 and essentially pure nitration grade benzene is removed overhead through line 101 and a nearly pure toluene stream is removed through line 102.

The following examples are illustrative of the present invention.

Example 1 A gas oil derived from Western Canadian Crude is fed to a steam cracker.

GAS OIL FEEDSTOCK INSPECTIONS API, ASTM distillation 29.6 I.B,P. 438 5% 530 571 50% 646 90% 732 F.B.P. 751

Wt. percent sulfur 0.70

In the cracker it is mixed with 160 1b. of steam per barrel of oil and maintained at 1430 F. for a residence time of 0.8 second. 45 wt. percent of the feed is converted to C3 hydrocarbons and lighter. A 22 wt. percent yield of naphtha based on feed is obtained. After caustic treat ment this has the inspections shown in column 1 of Table 1.

This naphtha is hydroned over a reduced nickel on kieselguhr catalyst in the form of 8-12 mesh granules. The conditions used are:

Reactor temperature 300 F.-maintained essentially constant by cooling.

Pressure 200 p.s.i.g.

Treat gas 1000 s.c.f./bbl. of reformer tail gas.

Space velocity 1 v./v./hr.

Hydrogen consumption 470 s.c.f./b.

The inspections on the first stage hydroned product are shown in column 2 of Table 1.

using diethylene glycol-water solvent and the extract is clay treated and distilled into essentially pure benzene and a bottom product consisting of toluene and higher boiling aromatics.

TABLE 1 Naphtha Product; Product Feed from from Benzene to 1st First Second Product Stage Stage Stage LV% Yield on Naplitlia Feed 100 l00 100 11.7

API 32.5 33. 0 33. 5 28. 6 Sulfur, ppm.-- l, 700 1, 700 2 Bromine No 130 82 6 0. 5 Diene No. 30 5 1 0.1 Exist Gurn l 5 1 1 (2) Accel. Guml Benzene, LV% Thionhene, ppm MON Clear +3 ce. TEL RON Clear +3 cc. TEL 15/5 Dist. I.B.P

Acid Wash Color Aeidity Copper Corrosion 1 iiillibited with 10 lb./MB (Active) Dupont No. 5. 2 l i 3 Pass.

Example 2 The results obtainable by employing the scheme illustrated in FIGURE 2 using the same feedstock and cracking conditions described in Example 1 are summarized below in Table 2.

TABLE 2 N aphtha Product 20G/400 F. 1110/200 F. Product Feed to from 1st N aphtha Fraction from 2nd Benzene 1st Stg. Stg. rgn 1st to 2nd Stg. Stg. Product LV% Yield on N aphtha Feed 100 100 50. 2 22.1 22. 6 12.3 API Gravity 33. 1 43. 3 44. O 28. 6 Sulfur, p p m 1,400 95 8 1 Bromine No 4 49 3 0. 5 Dieue No 4.1 2.1 0. 5 0. 1 Exist Gum, mg./100 ml 1 1.0 (2) Accel. Gum, Ing/100 ml l. 15/5 Dist. I.B.P

Thiopliene, p.p.rn Freezing Pt., C Acid Wash Color Acidity Copper Corrosion l Inlliibited with l0 ltr/MB (octane )Dupont No. 5.

2 Ni 3 See Table 1.

4 Value may be high due to olefin interference.

5 Pass.

This material is fed to the second stage hydroner where it is treated over Ia sulded cobalt molybdateon-alurnina catalyst of 5/64 diameter extrudate. The conditions used are:

The inspections of the second stage hydroned product are shown in column 3 of Table 1. This material is now depentanized and approximately 15 volume percent of pentanes are removed.

The remaining material is extracted in a Udex plant The foregoing description contains a limited number of embodiments of the present invention. lt will be understood, however, that numerous variations are possible without departing from the scope of this invention.`

What is claimed is:

1. A process for t-he separation of aromatics from a feedstock mixture containing, in a major proportion, said aromatics and mono-olefins, `and a minor proportion of gum-forming constituents, comprising the steps of:

(l) contacting said feedstock in a first hydrogenation stage with a supported nickel catalyst which, under operating conditions, contains a major proportion of nickel in elemental form, in the presence of a gas in which the reactant thereof consists essentially of hydrogen and at an elevated temperature, whereby at least a substantial proportion of the gumforming' constituents are destroyed;

(2) passing at least a portion of the product stream from said first stage to a second, 'olefin removing, hydrogenation stage and contacting said portion with a hydrogenation catalyst different from said supported nickel catalyst of said first stage in the presence of hydrogen and at an elevated temperature whereby at least a substantial proportion of the mono-olefins in said portion are converted into parains; and

(3) recovering the aromatics from the product stream of said second stage.

2. A process according to claim l, wherein a separate hydrogen feed is maintained to each of said hydrogenation stages.

3. A process for the separation of aromatics from a feedstock consisting of a raw steam cracked gasoline, comprising the steps of:

(l) distilling said raw steam cracked gasoline and recovering a light naphth-a fraction therefrom predominating in hydrocarbons, said hydrocarbons being predominantly aromatics and mono-olefins with a minor proportion of gum-forming constituents;

(2) passing said fraction to a first, hydrogenation stage and contacting said fraction with a supported nickel catalyst which, runder operating conditions, contains a major proportion of nickel in elemental form, in the presence of a gas in which the reactant thereof consists essentially of hydrogen and at an elevated temperature whereby at least a substantial proportion of said gum-forming constituents are destroyed without substantial hydrogenation of the aromatic and mono-olefin constituents;

(3) passing at least a portion of the product stream from said first stage to a second, olefin removing,

stage and contacting said portion with a hydrogenation catalyst different from said supported nickel catalyst of said first stage in the presence of hydrogen and at an elevated temperature to convert the said mono-olefins of said fraction to paraffins; and,

(4) thereafter recovering the aromatics from the product stream of said second stage.

4. A process for the separation of aromatics from a feedstock consisting of raw steam cracked gasoline containing predominantly aromatics and mono-olefins with a minor proportion of gum-forming constituents, comprising the steps of (l) contacting said feedstock with a supported nickel catalyst which, under operating conditions, contains a major proportion of nickel in elemental form, in the presence of a gas in which the reactant thereof consists essentially of hydrogen and at an elevated temperature, whereby at least a substantial proportion of the gum-forming constituents are destroyed;

(2) passing at least a portion of the product stream from said first stage to a second, olefin removing, hydrogenation stage and contacting said portion with a hydrogenation catalyst different from said supported nickel catalyst of said first stage in the presence of hydrogen and at an elevated temperature whereby at least a substantial proportion of the mono-olefins in said portion are converted into parafiins; and

(3) recovering the aromatics from the product stream of said second stage.

5. A process for the separation of aromatics from a feedstock mixture containing aromatics and mono-olefins in major proportions and gum-forming constituents in a minor proportion which comprises:

(l) contacting said feedstock in a first hydrogenation stage with a supported nickel catalyst which, under operating conditions, contains a major proportion of nickel in elemental form, in the presence of a gas in which the reactant thereof consists essentially of hydrogen and at a temperature in the range 250- 400 F., whereby at least a substantial proportion of the gum-forming constituents are destroyed,

(2) passing the product stream from said first stage to a second olefin removing hydrogenation stage and contacting said product stream in the presenceof hydrogen with a cobalt oxide-molybdenum oxide catalyst on a support, at a pressure in the range 400 to 1000 p.s.i.g., a temperature in the range S50-700 F., and at a `space velocity of 0.5 to4'.0 v./v./hr., whereby at least a substantial proportion of the mono-olefins in said product stream are converted into paraffins, and

(3) recovering the aromatics from the product stream of said second stage.Y

6. A process for the preparation of high purity aromatics which comprises:

(l) steam cracking a hydrocarbon feedstock boiling up to and including the gas oil boiling range,

(2) separating a naphtha fraction containing monoolefins and polyolefins from the steam cracked products,

(3) treating the naphtha fraction with hydrogen in contact with a catalyst' comprising reduced nickel on a support at temperatures of from about Z50-400 F. and pressures of about 150-300 p.s.i.g. to hydrogenate a major proportion of the polyolefins therein in a first hydrofining stage,

(4) treating the hydrofined naphtha product from the first hydrofining stage with hydrogen in contact with a metal oxide type hydrogenation catalyst at temperatures of about 550-700 F. and pressures of about 400-1000 p.s.i.g. in a second hydrofining stage to substantially saturate the mono-olefins therein,

(5) subjecting the hydrofined liquid product from the second hydrofining stage to solvent extraction.

7. The process as defined in claim 6 wherein the temperature in the first hydrofining stage is controlled by recycling hydrofined liquid product thereto.

8. The process as defined in claim 6 in which hydrofined liquid product from the second hydrofining stage is depentanized before extraction.

9. The process as dened in claim 6 in which the aromatics fraction extracted from the hydrofined liquid product from the second hydrofining stage is clay treated at 40G-600 F. to remove any traces of olefins or diolefins therefrom.

10. Process according to claim 6 in which the solvent is sulfolane.

11. Process according to claim 6 in which the solvent is a glycol-type solvent.

12. A process for the preparation of high purity benzene and toluene which comprises:

( l) steam cracking a hydrocarbon feedstock boiling up to and including the gas oil boiling range,

(2) separating a naphtha fraction containing monoolefins and polyolefins from the steam cracked products,

(3) treating the naphtha fraction `with hydrogen in contact with a reduced nickel` catalyst at temperatures of from about Z50-400 F. and pressures of about 150-300 p.s.i.g. to hydrogenate a major proportion of the polyolefins therein in a first hydrofining stage,

(4) fractionating the hydrofined naphtha product from the first hydrofining stage into a light fraction boiling up to F., a 140-200 F. aromatics-rich heart cut and a heavy 200-400" F. fraction,

(5) treating the 14C-200 F. fraction with hydrogen in contact with a metal oxide type hydrogenation catalyst at temperatures of about 55C-700 F. and pressures of from about 400-1000 p.s.i.g. in a second hydrofining stage to substantially saturate the mono-olefins therein,

(6) recovering a hydrofined liquid product having a maximum bromine number of about 5,

(7) subjecting said hydroned liquid product to solvent extraction to extract pure benzene and toluene therefrom.

13. A process for the preparation of high purity benzene and toluene which comprises:

(1) steam cracking a hydrocarbon feedstock boiling in the gas oil boiling range at temperature of at least about 1300 F.,

(2) separating a naphtha fraction from the steam cracked products,

(3) caustic washing said naphtha fraction to remove mercaptans therefrom,

(4) treating the caustic Washed naphtha fraction with hydrogen in contact with hydrogen in contact with a reduced nickel catalyst at temperatures of from about Z50-400 F. and pressures of about 150-300 p.s,i.g. to hydrogenate a major proportion of the polyolens therein in a first hydroining stage,

(5 fractionating the hydroined naphtha product from the first hydroning stage into a light fraction boiling up to 140 F., a 140-200 F. aromatics-rich heart cut and a heavy ZOO-400 F. fraction,

(6) treating the 140-200 F. fraction with hydrogen in contact with a metal oxide type hydrogenation catalyst at temperatures of about 550-700 F. and pressures of from about 400-1000 p.s.i.g. in a second hydroning stage to substantially saturate the mono-olens therein,

(7) subjecting the hydroned liquid product from the second hydroning stage to solvent extraction with a glycol-type solvent to extract pure benzene and toluene therefrom.

14. Process according to claim 13 in which said metal oxide type catalyst is sulfided.

References Cited UNITED STATES PATENTS DELBERT E. GANTZ, Primary Examiner.

CHARLES E. SPRESSER, Assistant Examiner. 

1. A PROCESS FOR THE SEPARATION OF AROMATICS FROM A FEEDSTOCK MIXTURE CONTAINING, IN A MAJOR PROPORTION, SAID AROMATICS AND MONO-OLEFINS, AND A MINOR PROPORTION OF GUM-FORMING CONSTITUENTS, COMPRISING THE STEPS OF: (1) CONTACTING SAID FEEDSTOCK IN A FIRST HYDROGENATION STAGE WITH A SUPPORTED NICKEL CATALYST WHICH, UNDER OPERATING CONDITIONS, CONTAINS A MAJOR PROPORTION OF NICKEL IN ELEMENTAL FORM, IN THE PRESENCE OF A GAS IN WHICH THE REACTANT THEREOF CONSISTS ESSENTIALLY OF HYDROGEN AND AT AN ELEVATED TEMPERATURE, WHEREBY AT LEAST A SUBSTANTIAL PROPORTION OF THE GUMFORMING CONSTITUENTS ARE DESTROYED; (2) PASSING AT LEAST A PORTION OF THE PRODUCT STREAM FROM SAID FIRST STAGE TO A SECOND, OLEFIN REMOVING, HYDROGENATION STAGE AND CONTACTING SAID PORTION WITH A HYDROGENATION CATALYST DIFFERENT FROM SAID SUPPORTED NICKEL CATALYST OF SAID FIRST STAGE IN THE PRESENCE OF HYDROGEN AND AT AN ELEVATED TEMPERATURE WHEREBY AT LEAST A SUBSTANTIAL PROPORTION OF THE MONO-OLEFINS IN SAID PORTION ARE CONVERTED INTO PARAFFINS; AND (3) RECOVERING THE AROMATICS FROM THE PRODUCT STREAM OF SAID SECOND STAGE. 